Dual mode reactor SMR integration

ABSTRACT

The present invention relates to systems and processes for producing syngas in steam methane reformer (SMR)-based plants, particularly to the use of a high space velocity, dual mode catalytic reactor to pre-reform plant feedstock. The dual mode reactor has the capability to operate in two modes: either without oxygen addition in a reforming mode or with oxygen addition in a partial oxidation-reforming mode. The dual mode reactor allows the syngas production rate of the plant to be manipulated without the added capital expense of a reheat coil and with reduced impact on export steam production.

CROSS-REFERENCE TO RELATED APPLICATION

The present application claims priority to U.S. provisional patentapplication Ser. No. 60/931,182, filed May 22, 2007 and claims priorityto U.S. provisional patent application Ser. No. 60/964,877, filed Aug.15, 2007; the entire contents of both U.S. provisional patentapplication Ser. No. 60/931,182 and U.S. provisional patent applicationSer. No. 60/964,877 are incorporated herein by reference in theirentirety.

TECHNICAL FIELD

The present invention relates to systems and processes for producingsyngas in steam methane reformer (SMR)-based plants, and moreparticularly to methods and systems that include high space velocity,dual mode catalytic reactors used to pre-reform feedstock and manipulatethe syngas production rates of such SMR-based plants.

BACKGROUND ART

Steam methane reformers (SMRs) are commonly used to produce a syngasincluding hydrogen and carbon monoxide from a gas feedstock such asnatural gas or refinery gas. The produced syngas can be furtherprocessed within the plant to yield various end products, includingpurified hydrogen, carbon monoxide, methanol and/or ammonia.Conventional adiabatic pre-reformers can be incorporated into SMR-basedsyngas plants for different reasons. For example, pre-reformers may beimplemented in order to: (i) reduce the content of ethane and heavierhydrocarbons in the feedstock to the SMR, thereby reducing thepropensity for carbon formation on the SMR catalyst and potentiallyenabling higher SMR feed temperatures or lower SMR feed steam-to-carbonmolar ratios, and/or (ii) increase the production rate of syngas orsyngas-derived products from the plant for a fixed SMR furnace duty.

Conventional adiabatic pre-reformers are catalytic reactors typicallycharged with pellet or shaped supported catalyst loaded with a high Nicontent. Typical pre-reformer gas hourly space velocities (GHSVs) basedon total feed at standard conditions (i.e. 60° F. and 1 atm) can rangefrom 15,000/hr to 25,000/hr. Typical pre-reformer feeds are steam andhydrocarbon mixtures with a steam-to-carbon molar ratio of 2 to 3,preheated to between 900° and 1200° F. Temperature change across thepre-reformer is dictated by the hydrocarbons in the feed. Whensignificant methane steam reforming occurs, temperatures decrease alongthe reactor due to the endothermic reaction.

In order to increase the production rate of the plant for a constant SMRfurnace duty, the pre-reformer effluent must be reheated against SMRflue gas, typically to 1100°-1200° F. Installation of this reheat coilcan be expensive, especially for existing plants, as the SMR flue gasduct must be opened and modified. Removing heat from the SMR flue gasthrough the reheat coil also reduces the amount of export steam that canbe produced in the downstream flue gas boiler and sold to customers.

U.S. Pat. No. 7,037,485 B1 to Drnevich et al. relates to the use of adual mode reactor for converting olefin containing gas for use as a feedfor a steam methane reformer. The first mode is a hydrogenation modewhile the second mode is a partial oxidation mode. Some pre-reforming isaccomplished in the second partial oxidation mode, but the capability tosubstantially manipulate the syngas production rate of the plant islimited, since the steam-to-carbon molar ratios in the reactor feed areto be maintained below 0.5, due to the presence of downstream sulfurremoval beds.

U.S. Pat. No. 6,335,474 B1 to Ostberg et al. relates to the use of anoble metal catalyst on a MgO and/or MgAl₂O₄ spinel carrier topre-reform hydrocarbon feedstocks containing oxygen impurities.

U.S. Patent Application No. 2005/0207970 A1 to Garg et al. considers thepre-reforming of natural gas over a nickel catalyst using oxygen in anamount less than that stoichiometrically required to partially oxidizeall ethane and heavier hydrocarbons in the natural gas to carbonmonoxide and hydrogen. For most natural gases, this translates intooxygen-to-natural gas molar ratios of less than 0.1, an amount thatwould not substantially increase the production rate of the SMR-basedsyngas plant.

SUMMARY OF THE INVENTION

The present invention relates to systems and processes for producingsyngas in steam methane reformer (SMR)-based plants, and moreparticularly to the use of a high space velocity, dual mode catalyticreactor to pre-reform plant feedstock. The dual mode reactor has thecapability to operate in two modes: either without oxygen addition in areforming mode or with oxygen addition in a partial oxidation-reformingmode. It is expected that the dual mode reactor will allow the syngasproduction rate of the plant to be manipulated without the added capitalexpense of a reheat coil and with reduced impact on export steamproduction.

As used herein, space velocity refers to gas hourly space velocity(GHSV), which is the ratio of the volumetric flow of the total reactorfeed at standard conditions of 60° F. and 1 atm to the catalyst volume.As used herein, high space velocity(ies) refers to GHSV(s) of greaterthan 30,000/hr.

BRIEF DESCRIPTION OF THE DRAWINGS

For a more complete understanding of the present invention and theadvantages thereof, reference should be made to the following DetailedDescription taken in conjunction with the accompanying drawings inwhich:

FIG. 1 is a schematic illustrating conventional steam methane reformer(SMR) technology employed in a hydrogen plant;

FIG. 2 is a schematic illustrating a standard SMR with a conventionaladiabatic pre-reformer (with some plant equipment shown in FIG. 1 havingbeen excluded for clarity);

FIG. 3 is a schematic illustrating the dual mode reactor of the presentinvention integrated with the standard SMR in accordance with apreferred embodiment of the present invention (with some plant equipmentshown in FIG. 1 having been excluded for clarity);

FIG. 4 is a schematic illustrating the dual mode reactor of the presentinvention integrated with the standard SMR in an alternative embodimentof the present invention (with some plant equipment shown in FIG. 1having been excluded for clarity);

FIG. 5 is a schematic illustrating a dual mode reactor of the presentinvention integrated with the standard SMR in yet another alternativeembodiment of the present invention which includes processing afeedstock partly comprised of refinery gas (with some plant equipmentshown in FIG. 1 having been excluded for clarity); and

FIG. 6 is a schematic illustrating the dual mode reactor of the presentinvention integrated with the standard SMR in yet another alternativeembodiment of the present invention (with some plant equipment shown inFIG. 1 having been excluded for clarity).

DETAILED DESCRIPTION

As discussed hereinabove, the present invention relates to systems andprocesses for producing syngas in steam methane reformer (SMR)-basedplants, and more particularly to the use of a high space velocity, dualmode catalytic reactor to pre-reform plant feedstock. The dual modereactor has the capability to operate in two modes: either withoutoxygen addition in a reforming mode or with oxygen addition in a partialoxidation-reforming mode. The dual mode reactor is expected to allow thesyngas production rate of the plant to be manipulated without the addedcapital expense of a reheat coil and with reduced impact on export steamproduction. As used herein, space velocity refers to gas hourly spacevelocity (GHSV), which is the ratio of the volumetric flow of the totalreactor feed at standard conditions of 60° F. and 1 atm to the catalystvolume. As used herein, high space velocity(ies) refers to GHSV(s) ofgreater than 30,000/hr.

Standard SMR

Referring now to FIG. 1, a schematic diagram of conventional steammethane reformer technology used as the basic process for producingintermediate to large quantities of hydrogen in hydrogen plant 600 foruse in refineries and other applications is illustrated.

More specifically, natural gas 1 is mixed with a small amount ofhydrogen product 2 to form stream 4 that is preheated in product heatrecovery system 135. The heated stream 6 is hydrotreated and sulfur isremoved in combined hydrotreater adsorber 130. The desulfurized feedstream 8 is mixed with superheated steam 11 to form stream 24. Steam 11is produced by superheating saturated steam 92 against flue gas 40 inheat recovery unit 115, which is also known as the reformer's convectionsection. The steam to carbon ratio in stream 24 can vary depending onthe design but normally is in the range of about 3/1. As also shown inFIG. 1, the natural gas-steam mixture 24 is further heated against fluegas 40, generating stream 26. Stream 26, typically at 900 to 1200° F.,is introduced into the reformer tubes 106 contained in steam methanereformer 100. The internal volume of the reformer tubes 104 are filledwith catalyst, usually composed of nickel compounds. The catalystpromotes the conversion of the natural gas-steam mixture to hydrogen andcarbon monoxide. Gas temperatures in the tube-side reformer typicallyrange from about 900° F. to about 1700° F. Tube-side gas temperaturesincrease from the reformer inlet to the exit. The maximum tube-side gastemperature, normally about 1600° F., is at the reformer exit. Both thesteam methane reforming reaction:CH₄+H₂O

CO+3H₂and the shift conversion reaction:CO+H₂O

CO₂+H₂take place within tube volume 104. The reformed gas exits reformer 100as stream 46, mainly including a mixture of hydrogen, carbon monoxide,carbon dioxide, water vapor and unreacted methane. Typical hydrogencontents of stream 46 may range from 46 to 50 mol %.

Stream 46 is cooled in process-gas heat-recovery system 135 against hotwater producing steam. After steam is generated, the still hot syngasleaves unit 135 as stream 48 and enters the shift conversion unit 125where the shift reaction is driven further to the right. The shiftconversion reaction is slightly exothermic and the unit(s) normallyoperates at temperatures ranging from about 400° F. to about 900° F. Inthis case stream 50, leaving the shift conversion reactor at about 800°F., is reintroduced to unit 135 where it is cooled against the feed gas4 and various streams containing water. Gas 51 exits process heatrecovery section 135 and enters knockout drum 142, in which condensedwater vapor is separated from the process stream. Gas 52 is furthercooled in unit 140 against cooling water and/or through the use offin-fan type air coolers. Cooled gas 53 enters knockout drum 144, inwhich additional condensed water vapor is separated from the processstream. Cooled and dried stream 54 is introduced into pressure swingadsorption (PSA) unit 145. The PSA produces hydrogen 56 at puritiesranging from about 99% to about 99.999% based on the system design. ThePSA hydrogen recovery can range from about 75% to about 95%. Theunrecovered hydrogen and any carbon monoxide, methane, water vapor,carbon dioxide and nitrogen present in stream 54 are purged from the PSAunit as tail gas 58. The tail gas is normally sent back to the reformerto be used as fuel.

Additional natural gas 32 and, for hydrogen plants with PSApurification, PSA tail gas 58 are burned with preheated air 30 inburners (not shown) to provide the heat to drive the reformingreactions. The burner exhausts into the “radiant” section of thereformer 102 where the heat generated through combustion is transferredby radiant and convective mechanisms to the surface of tubes 106. Heatfrom the tube surface is conducted to the interior of the tubes andtransferred to the process gas through convection. The tube walltemperature is a critical parameter influencing the life of the tubes.Excess temperatures can dramatically reduce the time between tubereplacements. The flue gas 40, leaving the radiant section attemperatures ranging from about 1600° F. to about 2000° F., enters theconvection section 115 where the contained sensible heat is used topreheat the natural gas-steam mixture as well as produce and superheatsteam. The flue gas 42 leaving the convection section 115 enters aninduced draft fan 120 which is used to maintain the radiant section ofthe reformer at a pressure slightly below atmospheric. Stream 44 is sentto a flue stack where it is vented to the atmosphere, normally attemperatures in excess of about 260° F.

Stream 60, a mixture of condensate and makeup boiler feedwater, isheated in unit 135, then de-aerated in unit 148. Steam 96 is commonlyused as a purge gas in the de-aerator. The de-aerated boiler feed wateris pumped in unit 155 to the pressure needed to provide superheatedsteam at sufficient pressure for mixing with natural gas to producestream 24 and/or high enough to provide superheated steam for export.Stream 66 is split into stream 68 and 70. Stream 68 is sent to unit 135where it is heated to near the boiling temperature. Stream 72 is thensplit into stream 74 and 76. Stream 74 is boiled in unit 135. Stream 70goes to unit 115 where it is heated to near the boiling temperature.Stream 80 is mixed with stream 76 to form stream 82 and then is splitinto streams 84 and 88 that go to units 135 and 115, respectively, to bevaporized. Saturated steam 86 and 90 from unit 115 and unit 135,respectively are mixed with stream 78 in saturated steam header 94. Mostof the steam is sent as stream 92 to be superheated in unit 115. A smallquantity 96 is sent to the deaerator 148. The superheated steam leavesunit 115 as stream 10 and is split into stream 11 for mixing with thenatural gas feed to the reformer and into stream 22 which can be sold,used to produce electricity, or used to provide heat to unit operationsassociated with a refinery or chemical plant operations.

Standard SMR with Conventional Adiabatic Pre-Reformer

FIG. 2 is a schematic of a basic SMR that includes a conventionaladiabatic pre-reformer. Conventional adiabatic pre-reformer 149 is acatalytic reactor typically charged with pellet or shaped supportedcatalyst loaded with a high Ni content. The reactor is sized such thatthe gas hourly space velocity (GHSV) based on total feed at standardconditions (i.e. 60° F. and 1 atm) is 15,000/hr to 25,000/hr.Conventional adiabatic pre-reformer 149 can be incorporated into steammethane reformer (SMR)-based syngas plants for different reasons. Forexample, a pre-reformer may be implemented in order to: (i) reduce thecontent of ethane and heavier hydrocarbons in the feedstock to the SMR,thereby reducing the propensity for carbon formation on the SMR catalystand potentially enabling higher SMR feed temperatures or lower SMR feedsteam-to-carbon molar ratios, and/or (ii) increase the production rateof syngas or syngas-derived products from the plant for a fixed reformerfurnace duty.

Stream 24, which includes heated, desulfurized natural gas 8 andsuperheated steam 11, is further heated against SMR flue gas in themixed feed preheat coil 162 to between about 900° and 1200° F. Resultingstream 26 a is fed to conventional adiabatic pre-reformer 149, wherehydrocarbon steam reforming and water gas shift reactions occur,producing hydrogen, carbon monoxide and carbon dioxide. Due to these netendothermic reactions, pre-reformer effluent 26 b emerges at a lowertemperature than stream 26 a with little or no remaining ethane andheavier hydrocarbons. The extent of these reactions is primarilydictated by the temperature of feed 26 a.

If pre-reformer effluent 26 b were directly fed to SMR 100 operating ata constant fired duty, the reduced stream temperature of stream 26 bwould prevent any increase in the plant 600 hydrogen production rate. Inorder to achieve an increase in the production rate of the plant 600,pre-reformer effluent 26 b must therefore be reheated against SMR fluegas 40 in reheat coil 160, typically to temperatures between about 1100°to 1200° F. The reheated temperature of stream 26 may be dictated bypiping metallurgy and/or by carbon formation concerns. Installation ofreheat coil 160 can be expensive, especially for existing plants, as theSMR flue gas duct 115 must be opened and modified. Removing heat fromthe SMR flue gas through reheat coil 160 also reduces the amount ofexport steam 22 that can be produced and sold to customers, because lessheat is available to the downstream boiler feed water heater 168, fluegas boiler 166 and steam superheater 164.

Dual Mode Reactor

Referring now to FIG. 3, a preferred embodiment of the dual mode reactorintegrated with the standard SMR in accordance with the presentinvention is shown. The dual mode reactor 150, which replaces theconventional adiabatic pre-reformer, is a high space velocity catalyticreactor that has the capability to operate in two modes: either withoutoxygen addition in a reforming mode or with oxygen addition in a partialoxidation-reforming mode. As discussed above, space velocity refers togas hourly space velocity (GHSV), which is the ratio of the volumetricflow of the total reactor feed at standard conditions of 60° F. and 1atm to the catalyst volume. High space velocity(ies) refers to GHSV(s)of greater than 30,000/hr. To-date, experimental data have confirmedthat space velocities of about 46,000/hr can be desirable for reactor150 performance. However, catalysts used in reactor 150 at even higherspace velocities could be advantageous.

In any embodiment of the present invention, dual mode catalytic reactor150 preferably contains a Group VIII catalyst supported on a metallicmonolith. That is, dual mode catalytic reactor 150 contains a catalystwhich is preferably a metallic monolith coated with a catalytic layerthat contains platinum, rhodium, palladium, nickel, ruthenium, or acombination of these metals. The structure of the monolith can bereticulated foam, honeycomb or a corrugated foil wound in a spiralconfiguration. It is believed that the metallic monolith supportedcatalyst has better performance than other supported catalyst in that ithas better heat conductivity and a more uniform temperature profile thanother catalyst forms. However, catalyst coated beads or ceramicmonoliths in the form of a reticulated foam or honeycomb structure areother options. One viable catalyst in the form of a monolith iscommercially available from Sud-Chemie, Inc. of Louisville, Ky., USA asPC-POX 1 on FeCrAlY. Similar catalysts from other suppliers may be used.

In partial oxidation-reforming mode, the effluent 26 a from the mixedfeed preheater 162 typically at 900 to 1200° F. is split into streams 26c and 26 e. Stream 26 e, which is preferably 10 to 75% and morepreferably 20 to 50% of stream 26 a, is fed to dual mode catalyticreactor 150. Oxygen-containing stream 200, preferably greater than 99.0%pure oxygen, is also introduced into reactor 150, where it reacts withstream 26 e in a combination of partial oxidation, complete oxidation,hydrocarbon steam reforming and water gas shift reactions. It ispreferred that sufficient oxygen is added such that the net reaction isexothermic and stream 26 f emerges at a higher temperature than stream26 e, typically ranging from 1100 to 1800° F. Reactor effluent 26 f isquenched with the lower temperature bypassed stream 26 c to form stream26 d, which is fed to SMR 100. Though not shown, reactor effluent 26 fcould be additionally quenched by adding a portion of boiler feed water(e.g. stream 80), steam (e.g. stream 86), superheated steam (i.e. 11),preheated hydrocarbon feedstock (i.e. 8) and/or mixed feed (i.e. stream24). By using these additional quench streams, the steam-to-carbon molarratio of the feed to reactor 150 (i.e. stream 26 e) can be maintained ata different value than that of the feed to SMR 100 (i.e. stream 26 d).

Unlike conventional adiabatic pre-reformer 149 of FIG. 2, the extent ofhydrocarbon reforming within dual mode catalytic reactor 150 is notprimarily dictated by the temperature of feed 26 a. Instead, the amountof hydrocarbon reforming can be adjusted by manipulating the flow ofoxygen-containing stream 200. More oxygen results in higher reactor exittemperatures and more hydrocarbon reforming. Both the selected oxygenflow and the amount of mixed feed 26 a sent to the reactor 150 as stream26 e dictate the total achievable increase in the plant 600 hydrogenproduction rate. Bypass stream 26 c, oxygen-containing stream 200, andany other reactor 150 effluent quench streams may be manipulated toobtain the desired hydrogen production rate while maintaining stream 26d below some maximum allowable temperature dictated by piping metallurgyand/or by carbon formation concerns.

Even though some of the mixed feed 26 a is bypassed around the dual modecatalytic reactor 150 as stream 26 c, more total hydrocarbon reformingcan be achieved relative to the conventional adiabatic pre-reformer ofFIG. 2, since the reactor 150 of FIG. 3 generally operates at higherexit temperatures due to the addition of oxygen. Since reactor 150 needonly be sized to process a portion of stream 26 a and since the designspace velocity for reactor 150 will be higher than the conventionalpre-reformer, a relatively smaller reactor vessel and less catalyst isrequired to achieve larger hydrogen production rates. Smaller reactorsizes may also allow for the use of parallel trains of reactor 150,which could increase reliability and decrease downtime for catalystchange-outs.

Unlike the conventional adiabatic pre-reformer of FIG. 2, the dual modecatalytic reactor 150 does not require a reheat coil in order toincrease hydrogen production rates. Instead, heat is generated in-situby exothermic reactions involving oxygen. Besides further reducingcapital costs, elimination of the reheat coil can also mitigate thereduction of export steam.

If additional hydrogen production is not required, injection ofoxygen-containing stream 200 can be terminated and the dual modecatalytic reactor 150 can be operated in reforming mode. Oxygen flow isterminated by closing the appropriate valves. Reactor bypass flow 26 ccould be maintained at the partial oxidation-reforming mode setpoint orcould optionally be reduced to zero by closing the appropriate valves.Reforming mode may be entered when the value of additional hydrogen isinsufficient to justify the additional cost of oxygen. Though noadditional hydrogen production from plant 600 will be achieved inreforming mode, hydrocarbon steam reforming and water gas shiftreactions will still occur and the ethane and heavier hydrocarboncontent of stream 26 d fed to SMR 100 will still be reduced. As usedherein, heavier hydrocarbons refers to all hydrocarbons containing twoor more carbon atoms.

Table 1 presents the simulated performance of hydrogen plant 600 forfive cases: (a) the base case described in FIG. 1, (b) the base caseretrofitted with a conventional adiabatic prereformer from FIG. 2, (c)the base case retrofitted with a dual mode reactor operating inreforming mode from FIG. 3, (d) the base case retrofitted with a dualmode reactor operating in partial oxidation-reforming mode from FIG. 3,and (e) case (d) with increased oxygen usage. Thus, cases (c) to (e)represent the same dual mode reactor system operating with varyinglevels of oxygen addition. All cases presume a desulfurized natural gasfeedstock, a constant SMR 100 fired duty of about 697.7 MMBTU/hr LHV anda constant mixed feed 24 steam-to-carbon ratio of 2.8. As used herein,“MMBTU/hr LHV” refers to the contained energy flow rate of the stream inmillions of BTUs per hour on a lower heating value basis.

For case (b), the conventional adiabatic pre-reformer 149 is presumed toachieve equilibrium at the simulated reactor exit temperature of 898° F.The pre-reformed feedstock 26 b is routed to the reheat coil 160, whereit emerges as stream 26 at 1200° F. The resulting simulated plant 600hydrogen production rate is 110 MMSCFD, a 10% increase over base case(a). As used herein, “MMSCFD” refers to the standard volumetric flowrate of the stream in millions of standard cubic feet per day at 1 atmand 60° F.

In simulated cases (d) and (e), both reactor bypass 26 c and a portionof mixed feed 24 are used to quench reactor 150 effluent 26 f,maintaining stream 26 d at a maximum temperature of 1150° F. Thesimulations indicate that case (d) uses 144 tons per day (tpd) of 99.9%pure oxygen in stream 200 to achieve a reactor 150 exit temperature of1400° F. and a plant 600 hydrogen production rate of 110.4 MMSCFD, whilecase (e) uses 222 tpd of 99.9% pure oxygen in stream 200 to achieve areactor exit temperature of 1600° F. and a plant 600 hydrogen productionrate of 115.4 MMSCFD. In the absence of the reheat coil, case (c)produces the same amount of hydrogen as the case (a) basis, namely 100MMSCFD. Cases (c) through (e) demonstrate that the stream 200 oxygen andthe reactor bypass 26 c can be manipulated in real time during plant 600operation to achieve variable hydrogen production rates from 100 to115.4 MMSCFD while using the same dual mode reactor 150. Stream 26 d SMRfeed can simultaneously be maintained at or below a maximum temperature.

Due to the bypassed flow and a larger GHSV, the dual mode reactor ofcases (c) to (e) requires almost 90% less catalyst (i.e. 36 vs. 347 ft3)than the conventional adiabatic pre-reformer of case (b). The case (d)and (e) simulations demonstrate that this smaller reactor 150 can stillachieve 0.4 and 5.4 MMSCFD more hydrogen production from plant 600,respectively, than the much larger conventional adiabatic pre-reformer149. Additionally, the retrofit installation of reheat coil 160 is notrequired. Both the smaller reactor vessel and the elimination of areheat coil retrofit should reduce capital costs.

Finally, for the case (b) adiabatic pre-reformer, the increase in plant600 hydrogen production rate comes at the expense of an estimated 24%reduction in export steam sales (i.e. 156 to 119 thousand pounds perhour or kpph). However, for the simulated dual mode reactor cases (d)and (e), even more hydrogen is produced with a net 4 to 9% increase inexport steam sales.

TABLE 1 Simulated comparison of present invention vs. conventionalalternative Case ID D E B C Dual Mode- Dual Mode- A Conventional DualMode- POX/Reforming- POX/Reforming- Case Descriptors Base Pre-reformerReforming Less O2 More O2 Reactor 149 or 150 Total Feed Flow, kscfh 69401803 1803 1803 Reactor 149 or 150 GHSV, 1/hr 20000 50000 50000 50000Reactor 149 or 150 Catalyst Volume, ft 3 347 36 36 36 Reactor 149 or 150Exit T, F 898 928 1400 1600 SMR Feed T, F 1120 1200 928 1150 1150Steam/Carbon to Reactor 149 or 150 2.8 2.8 2.8 2.8 Product H2, mmscfd100.0 110.0 100.0 110.4 115.4 SMR Fired Duty, mmbtu/hr LHV 697.7 697.6697.8 697.7 697.6 NG Feed, kscfh 1620 1787 1619 1821 1919 NG Fuel, kscfh124 55 125 95 84 Export Steam, kpph 156 119 156 163 170 O2, tpd 0 0 0144 222

While a reheat coil analogous to coil 160 in FIG. 2 could conceivably beused in conjunction with the FIG. 3 embodiment, it is not preferredgiven the additional expense and given its severe underutilization whileoperating in the partial oxidation-reforming mode. While such a reheatcoil would provide increased plant 600 output during operation inreforming mode, provisions would need to be made to avoid overheatingthe coil in partial oxidation-reforming mode. If a pre-existing reheatcoil were present at the syngas-producing plant, it could be integratedinto the present invention.

FIG. 4 shows an alternative embodiment of the present invention in whichthe dual mode reactor is integrated with the standard SMR such thatthere is no bypass around reactor 150. Reactor 150 of FIG. 4 would beoperated at lower temperatures (e.g. 1000°-1300° F.) than reactor 150 ofFIG. 3 for two reasons. First, all feedstock undergoes reaction in theembodiment shown in FIG. 4, and so a lower reactor 150 exit temperatureis required for a targeted amount of total hydrocarbon reforming.Second, in the embodiment shown in FIG. 4, reactor 150 exit temperatureswill be at least partly dictated by piping metallurgy and/or by carbonformation concerns, since there is no quenching from a reactor bypassstream. On the one hand, some disadvantages of the embodiment shown inFIG. 4 relative to the embodiment shown in FIG. 3 could include: (i)significantly larger reactor 150 feed flows resulting in the need forlarger vessel sizes and larger catalyst requirements and/or (ii) lessflexibility due to fewer process variables that can be manipulated. Onthe other hand, however, some advantages of the FIG. 4 embodimentrelative to the embodiment illustrated in FIG. 3 could include: (i) morecomplete elimination of ethane and heavier hydrocarbons since no mixedfeed is bypassed around reactor 150, (ii) lower catalyst temperatures,which could lengthen catalyst life, and/or (iii) elimination of any hightemperature valving required for the reactor bypass.

FIG. 5 shows the preferred embodiment of FIG. 3 in combination with anembodiment of U.S. Pat. No. 7,037,485 B1. The entire contents of U.S.Pat. No. 7,037,485 B1 are incorporated herein by reference. Morespecifically and as shown in FIG. 5, raw refinery gas containing olefins8 b is processed as taught by U.S. Pat. No. 7,037,485 B1, resulting in adesulfurized and heated refinery gas feedstock with reduced olefincontent 8 c. Stream 8 c is combined with desulfurized and heated naturalgas 8, resulting in stream 8 a, which is ultimately fed to dual modereactor 150.

Raw refinery gas 8 b is compressed in one or more intercooled compressorstages 500 and passed through an impurities removal bed 501 which mayremove sulfur species, metals, halides and other impurities of concern.The refinery gas is then heated against condensing steam in preheater502.

When dual mode reactor 503 from U.S. Pat. No. 7,037,485 is operating inhydrogenation mode, steam 552 may optionally be added to suppress carbonformation. In this mode, olefin hydrogenation reactions dominate reactor503, and the reactor effluent has a reduced olefin content.

When dual mode reactor 503 from U.S. Pat. No. 7,037,485 is operating inoxidation mode, steam 552 and oxygen-containing gas 551 are added. Inthis mode, partial oxidation, complete oxidation, hydrocarbon steamreforming and water gas shift reactions occur within reactor 503, andthe reactor effluent has both a reduced olefin content and an increasedcontent of carbon monoxide plus carbon dioxide plus methane. In eithermode, the effluent from reactor 503 emerges at a higher temperature thanthe inlet due to net exothermic reactions. Thus, the effluent is cooledin a boiler 504 and passed through a zinc oxide bed 505 to reduce sulfurto levels acceptable to the SMR 100 catalyst (e.g. <0.1 ppmv on a drybasis).

FIG. 5 illustrates that the hydrocarbon feed to the dual mode reactor150 need not be natural gas, but can be any hydrocarbon, includingrefinery gas or mixtures of refinery gas and natural gas. Otherapproaches for managing and processing refinery gases than thatdescribed in U.S. Pat. No. 7,037,485 may be used. It should be noted,however, that all hydrocarbons in the feedstock to the high spacevelocity, dual mode reactor in accordance with the embodiments of thepresent invention must be desulfurized to an acceptable level for theSMR 100 catalyst, preferably less than 1 ppmv and more preferably lessthan 0.1 ppmv on a dry basis, prior to being introduced into dual modereactor 150.

FIG. 6 shows an alternative embodiment of the present invention in whichthe dual mode reactor is integrated with the standard SMR such that thereactor 150 bypass stream 26 c is processed through reactor 150A, whichis either a conventional adiabatic pre-reformer or a dual mode reactoroperating without oxygen addition in reforming mode only. Hydrocarbonsteam reforming and water gas shift reactions occur within reactor 150A,producing hydrogen, carbon monoxide and carbon dioxide. Due to these netendothermic reactions, reactor 150A effluent 26 g emerges at a lowertemperature than stream 26 c with little or no remaining ethane andheavier hydrocarbons. Reactor 150A effluent 26 g also emerges at a lowertemperature than reactor 150 effluent 26 f. As such, reactor 150effluent 26 f is quenched by reactor 150A effluent 26 g, yielding astream 26 d which is fed to SMR 100. As previously described, reactoreffluent 26 f could be additionally quenched by adding a portion ofboiler feed water (e.g. stream 80), steam (e.g. stream 86), superheatedsteam (i.e. 11), preheated hydrocarbon feedstock (i.e. 8) and/or mixedfeed (i.e. stream 24). The resulting temperature of stream 26 d may bedictated by piping metallurgy and/or by carbon formation concerns.

On the one hand, some advantages of the embodiment shown in FIG. 6relative to the embodiment shown in FIG. 3 could include: (i) greaterreduction of total ethane and heavier hydrocarbons in stream 26 d and(ii) more effective use of stream 26 g as a quench for reactor 150effluent 26 f due to its lower temperature than stream 26 c. On theother hand, a disadvantage of the embodiment shown in FIG. 6 relative tothe embodiment shown in FIG. 3 could include greater capital costassociated with the reactor 150A and its associated catalyst.

While the above figures have described the use of high space velocity,dual mode catalytic reactors in a SMR-based hydrogen plant 600, otherSMR-based syngas production plants can be envisioned. For example andwhile not be construed as limiting, the produced syngas can be furtherprocessed within the plant to yield various end products, includingpurified hydrogen, carbon monoxide, methanol and/or ammonia.

Although the invention has been described in detail with reference tocertain preferred embodiments, those skilled in the art will recognizethat there are other embodiments within the spirit and the scope of theclaims.

1. A steam methane reforming method, the method comprising: introducinga first portion of a feedstock stream comprising steam and desulfurizedhydrocarbons into a high space velocity catalytic reactor configured fordual mode operation, one of the dual modes of operation being operationwithout oxygen addition in a reforming mode and another of the dualmodes of operation being operation with oxygen addition in a partialoxidation-reforming mode, wherein the high space velocity catalyticreactor is disposed downstream of a mixed feed preheater operating at atemperature range of about 900 to 1200° F., and a second portion of saidfeedstock stream bypass said high space velocity catalytic reactor priorto introducing the second portion of the feedstock into a steam methanereforming unit; reacting the first portion of said feedstock in the highspace velocity catalytic reactor such that a pre-reformed gas isproduced in the high space velocity catalytic reactor, the pre-reformedgas comprising reduced amounts of heavier hydrocarbons relative to thefeedstock; withdrawing the pre-reformed gas from the high space velocitycatalytic reactor; introducing the pre-reformed gas directly into saidsteam methane reforming unit; and reacting the pre-reformed gas in thesteam methane reforming unit to produce a syngas.
 2. The method of claim1, wherein high space velocity catalytic reactor contains a catalystsupported on a metallic monolith therein.
 3. The method of claim 2,wherein the metallic monolith is coated with a catalytic layercomprising one or more Group VIII metals.
 4. The method of claim 3,wherein the Group VIII catalyst comprises platinum, rhodium, palladium,nickel, ruthenium or mixtures thereof.
 5. The method of claim 2, whereinthe monolith is reticulated foam, honeycomb or a corrugated foil woundin a spiral configuration.
 6. The method of claim 1, wherein the syngasis further processed to produce a purified hydrogen stream.
 7. Themethod of claim 6, wherein the amount of syngas produced is adjustablebased on adjusting the amount of oxygen addition into the high spacevelocity, dual mode catalytic reactor during the partialoxidation-reforming mode.
 8. The method of claim 1, wherein the highspace velocity of the high space velocity, dual mode catalytic reactoris greater than 30,000/hr.
 9. The method of claim 1, wherein thepre-reformed gas is not subjected to heating after withdrawing thepre-reformed gas from the high space velocity, dual mode catalyticreactor and prior to introduction into the SMR unit.
 10. The method ofclaim 1, wherein a temperature of the pre-reformed gas being withdrawnfrom the high space velocity, dual mode catalytic reactor duringoperation in the partial oxidation-reforming mode is equal to or greaterthan a temperature of the feedstock being introduced into the high spacevelocity, dual mode catalytic reactor.
 11. The method of claim 10,wherein the temperature of the pre-reformed gas being withdrawn from thehigh space velocity, dual mode catalytic reactor is less than about1800° F.
 12. The method of claim 1, wherein the amount of syngasproduced is adjustable based on adjusting the amount of oxygen additioninto the high space velocity, dual mode catalytic reactor during thepartial oxidation-reforming mode.
 13. The method of claim 1, wherein theamount of syngas produced is adjustable based on adjusting the amount offeedstock that is bypassed around the high space velocity, dual modecatalytic reactor.
 14. The method of claim 13, wherein the amount ofsyngas produced is further adjustable based on the amount of oxygenaddition into the high space velocity, dual mode catalytic reactorduring the partial oxidation-reforming mode.
 15. The method of claim 1,wherein the pre-reformed gas withdrawn from the high space velocitycatalytic reactor is optionally combined with a portion of the feedstockstream and introduced into the steam methane reformer.